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Process Engineering Guide: GBHE-PEG-RXT-810

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Process Engineering Guide: Gas - Liquid Reactors CONTENTS SECTION 0 INTRODUCTION/PURPOSE 3 1 SCOPE 3 2 FIELD OF APPLICATION 3 3 DEFINITIONS 3 4 PRELIMINARY CONSIDERATIONS 3 4.1 Preliminary Equipment Selection 4 4.2 Equipment for Low Viscosity Liquids 4 4.3 Equipment for High Viscosity Liquids 7 5 REACTOR DESIGN 8 6 ESSENTIAL THEORY 8

6.1 Rate and Yield Determining Steps 8 6.2 Chemical and Physical Rates 9 6.3 Modification for Exothermic and Complex Reactions 13 6.4 Preliminary Selection of Reactor Type 14



7 EXPERIMENTAL DETERMINATION OF REGIME 15 7.1 Direct Measurement of Reaction Kinetics 16 7.2 Laboratory Gas-Liquid Reactor Experiments 17 8 EQUILIBRIUM AND DIFFUSIVITY DATA SOURCES 19 9 OVERALL EFFECTS 20 9.1 Liquid Flow Patterns 20 9.2 Scale of Mixing 21 9.3 Gas Flow Pattern : Mean Driving Force for Mass Transfer 21 9.4 Gas-Liquid Reactor Modeling 22 9.5 Heat Transfer 23 9.6 Materials of Construction 24 9.7 Foaming 24 10 FINAL CHOICE OF REACTOR TYPE 24 11 SCALE-UP AND SPECIFICATION OF GAS-LIQUID

REACTORS 26 11.1 Bubble Columns 26 11.2 Packed Columns 26 11.3 Trickle Beds 26 11.4 Plate or Tray Columns 26 11.5 Spray Columns 27 11.6 Wiped Film 27 11.7 Spinning Film Reactors 27 11.8 Stirred Vessels 27 11.9 Plunging Jet 27 11.10 Surface Aerator 27 11.11 Static Mixers 27 11.12 Ejectors, Venturis and Orifice Plates 28 11.13 3-Phase Fluidized Bed 28



12 BIBLIOGRAPHY 28 TABLES 1 REGIMES OF GAS-LIQUID MASS TRANSFER WITH

ISOTHERMAL CHEMICAL REACTION 12 2 REGIMES OF GAS-LIQUID MASS TRANSFER IGNORING

LARGE EXOTHERMS OR OTHER COMPLICATIONS 15 3 COMPARATIVE MASS TRANSFER PERFORMANCE OF

CONTACTING DEVICES 18 4 COMPARATIVE MASS TRANSFER DATA 24 5 CHOICE OF GAS-LIQUID REACTOR TYPE 25 FIGURES 1 RATE AND YIELD DETERMINING STEPS 8 2 ENHANCEMENT FACTOR vs HATTA NUMBER 11 3 ENHANCEMENT FACTOR vs HATTA NUMBER : EFFECT

OF THERMAL & OTHER FACTORS 14 4 REACTORS FOR LIQUID-PHASE KINETICS

MEASUREMENT 17 5 EXPERIMENTS TO DETERMINE THE OPERATING

REGIME 19 6 EXPERIMENTS DETERMINE THE OPERATING REGIME

WHERE A SOLID CATALYST IS INVOLVED 20 7 THE MIXED ZONES IN LOOPS' MODEL FOR STIRRED

REACTORS 23



DOCUMENTS REFERRED TO IN THIS PROCESS ENGINEERING GUIDE 30



0 INTRODUCTION/PURPOSE This Guide is one in a series on Reactor Technology produced by GBH Enterprises. Absorption of gases into liquids, followed by reaction in the liquid phase, forms an important category of reactions of relevance to GBHE. Among these are oxidation, hydrogenation, sulfonation, halogenation and carbonylation. Sometimes the reactions are assisted by a catalyst, either dissolved or suspended in the liquid phase. 1 SCOPE This Guide provides a basic introduction to the theory of gas-liquid reactions. It deals with the microscopic scale phenomena of absorption, mass transfer and chemical reaction and indicates how these are integrated with fluid flow and heat transfer to provide predictions of reactor performance. Some attention is given to the experimental determination of the relevant physicochemical parameters. Finally, guidance is offered on the selection of appropriate gas-liquid reactors for a given system. 2 FIELD OF APPLICATION This Guide applies to process engineers in GBH Enterprises worldwide. 3 DEFINITIONS For the purposes of this Guide, no specific definitions apply. With the exception of terms used as proper nouns or titles, those terms with initial capital letters which appear in this document and are not defined above are defined in the Glossary of Engineering Terms.



4 PRELIMINARY CONSIDERATIONS Reactions between gases and liquids are common in the chemical industry. Typical examples are: (a) Oxidation, hydrogenation, sulfonation, halogenation, carbonation, etc., (b) Fermentation, effluent aeration, (c) Polymerization, polycondensation, (d) Gas absorption (CO2, H2S, SO2, Cl2, NOX, etc.). In all these cases the reactant gas is first absorbed into the liquid and the reaction then occurs in the liquid phase. 4.1 Preliminary Equipment Selection The most commonly used equipment is either: (a) A liquid phase reactor (a vessel) into which the gas is introduced, or (b) A gas absorber in which the reaction occurs. These devices are not always successful or optimal. See Clause 10 for the appropriate design methods. Newer devices, e.g. gas-liquid static mixers, are expanding the range of operating regimes. The reactor type is dictated by the following: (1) Viscosity

Turbulent mixing, which is used whenever possible, demands different equipment from laminar mixing. The selection is thus divided into equipment for low and high viscosity fluids.



Note: The addition of a solvent can reduce the viscosity. (2) Relative Rates of Reaction and Gas-Liquid Mass Transfer

From a knowledge of the rate of reaction and the mass transfer rate, it is possible to determine which family of devices will be appropriate.

(3) Flow Pattern Requirements

The overall flow pattern requirements (i.e. plug flow or back mixed) narrow down the choice within a family.

(4) Residence Time Requirements

The residence time requirements are based on the previously determined rate-controlling step and also influence the selection.

(5) Other Constraints

The presence of solid particles, the need for heat transfer and the materials of construction will impose constraints on the selection process.

4.2 Equipment for Low Viscosity Liquids

For low liquid viscosities (say less than about 200 cP [0.2 Ns/m2]), turbulence can be used to obtain good mixing, high interfacial area and high heat and mass transfer coefficients. Degassing can be done by gravity. Suitable devices include:

(a) Simple Bubble Columns

The gas is sparged into the liquid at the bottom of the contactor.

Well established in plants and much design information available.



(b) Plate Columns The gas is redispersed and disengaged at each plate up the column.


(c) Packed Column & Trickle Bed

Gas and liquid film flow can be either co-current or counter-current.


(d) Mechanically Agitated Vessels

An impeller rotates in a tank (usually baffled) to give enhanced rates of mass transfer and heat transfer.


(e) Recirculating Bubble Columns

Gas is sparged into a 'riser' and disengaged at the surface. The liquid recirculation is driven by the voidage difference.

Established on plants, but only scant design information available.



(f) Surface Aerators

Agitators at the liquid surface entrain gas from the head space into the liquid.

Established on plants, but only scant design information available.

(g) In-line Static Mixers

Mixing energy is derived from the flow itself.

Newer devices; few plant installations, but fair small-scale information available.

(h) Jet Ejectors

The gas is sucked into the liquid flow in a tube and dispersed by intense turbulence.

Newer devices; few plant installations, but fair small-scale information available.

(j) Plunging Jets

Gas is entrained into the liquid surface by a liquid jet. Newer devices; few plant installations, but fair small-scale information available.



(k) Three Phase Fluidized Beds

The presence of the solid phase helps to promote good contact between the gas and liquid phases.

Newer devices; few plant installations, but fair small-scale information available. (l) Spinning Cone The "turbulent" film is generated centrifugally and gives high mass and heat fluxes. Newer devices; few plant installations, but fair small-scale information available.

4.3 Equipment for High Viscosity Liquids For high viscosity liquids, turbulence cannot be achieved practically and hence a laminar mechanism (cutting and folding) has to be used to incorporate the gas. Highly viscous systems are reluctant to degas under gravity, so degassing is best done in a thin-film device. Suitable devices include: (a) Dough Mixers Dough mixers (or dough beaters) incorporate the bubbles into the liquid surface. Widely used, but only scant design information is available.



(b) In-line Dynamic Mixers These are capable of generating high shear rates suitable for producing a high gas content. Widely used, but only scant design information is available. (c) Scraped-film Devices A thin film of liquid is generated by a blade moving near a surface. Widely used, but only scant design information is available. (d) Centrifugal Devices In these, e.g. a rotating disc contactor, a thin film is generated on the disc by centrifugal force. Widely used, but only scant design information is available. (e) Mechanically Agitated Vessel The bubbles are incorporated from the surface or dispersed by a separate high shear device in the base of the vessel. Widely used, but only scant design information is available.



The reaction conditions required will set the priorities for design. For example, in effluent aeration or many fermentations, the systems are dilute and the reactions are slow; so large vessels are used and energy efficiency is important but the mass transfer intensity requirements are modest. In chlorination’s and sulfonation’s, the reactions are generally fast and the gases quite soluble so high intensity mass and heat transfer are essential but contact times may be slow. With many oxidations and hydrogenations the gases are less soluble and longer contact time or gas recycling is needed; often the mixing pattern and solid suspension are also important. For food batters and creams the rheological limitations overrule, and control of the local shear rate and temperature is all important. 5 REACTOR DESIGN Ideally the reactor design involves a quantitative idea of both the chemical kinetics and the mixing and mass transfer. Analysis of the chemistry can be difficult or impossible and mass transfer prediction is uncertain because of the effects of bubble coalescence and flow pattern. Design from first principles is therefore not normally possible. Small scale experiments should be carried out to determine the relevant rates and other necessary reaction characteristics such as the heat of reaction, the degree of foaming and bubble coalescence, which are then scaled-up to the final plant design. If time and techniques permit, this experimentation also guides the choice of reactor type and indicates whether mathematical modeling should be carried out and if so, to what complexity. Generally, the combination of the difficulties associated with the analysis of the chemistry and the uncertainties in the gas-liquid fluid dynamics renders the scale-up of gas-liquid reactors a thorny procedure. Experiments should be carried out at the final plant conditions (pressure, temperature, catalyst concentration). Because of the difficulty and cost of experiments, allowances for the uncertainly be should made in scaling-up. Methods of calculation for both traditional and newer devices are presented in Clauses 6 to 9. They also give further guidance on the selection, experimentation and scale-up for gas-liquid reactions.



6 ESSENTIAL THEORY 6.1 Rate and Yield Determining Steps Most industrial reactions are multi-step reactions and therefore their description is generally simplified by focusing on the rate determining step. For parallel reactions, the yield is the outcome of competition between two or more simultaneous rates. The example shown in Figure 1, based on the oxidation of cyclohexane, illustrates the rate and yield determining steps: FIGURE 1 RATE AND YIELD DETERMINING STEPS

C is the desired product; reaction steps (1), (2) and (3) are first order in dissolved gas (oxygen), steps (4) and (5) are second order. The overall Rate is determined by step (1). The Yield of C is determined by two ratios, namely: Rate (2) Rate (4) i.e. the dissolved oxygen ratio, and

Rate (3) + Rate (5)

Rate (2)



For single phase reactions, consult GBHE-PEG-RXT-809. In gas-liquid reactions in which one or more reactants have to transfer from the gas phase to the liquid, the rate of this transfer has to be taken into account since it often modifies or controls the overall rate of a chemical step and possibly the rate or yield determining step. The theory and the implications are covered in 6.2 and 6.3; thus often mass transfer limitations are minimized to give high rates, but they can be useful to manipulate yields in complex reactions. 6.2 Chemical and Physical Rates For simple reactions the theory is well covered in several works (e.g. see Ref. [2], Ref. [14] and Ref.[15]) and is summarized below using the film theory. Note: Other theories can be used, but they result in almost identical predictions. The diffusion of a reactant through the gas to the interface and thence through the liquid film into the bulk liquid with simultaneous reaction is covered. Additional complications such as: (a) Multi-component diffusion; (b) Escape of volatile products; (c) Heats of reaction; (d) Concurrent reactions. Require the solution of more complex basic equations. The treatment centers around the relative rates of gas-liquid mass transfer and reaction. The reaction rate controls the overall rate if it is very slow; often, however, it is fast, so that the overall rate is mass transfer controlled. Very fast reactions can influence the diffusion process, causing enhancement of mass transfer above the purely physical rate. This enhancement is a function of the reaction rate. In extreme cases (so-called "instantaneous" reactions) mass transfer again controls the overall rate of the process. Consider a reaction between n moles of a dissolving gas A and m moles of a liquid phase reactant B. Assume negligible diffusion resistance in the gas phase.



nA + mB Products Let q = m/n Let CA and CB represent the molar concentrations of A and B respectively in the liquid. The rate of reaction of A is then described by:

Where: n and m are the orders of reaction in A and B, and

e is the liquid hold-up fraction.

Taking A as the limiting reactant, "reaction time” tR is defined by:

The mass transfer of A into the liquid is described by:

Where: k L= mass transfer coefficient a = interfacial area per unit reactor volume C*A = interface concentration of A



A mass transfer "diffusion time" tD may be defined as:

Where: DLA = diffusivity of A in the liquid. If a fast reaction occurs near the interface (within the "diffusion film") it will enhance the mass transfer rate and equation (2) becomes:

Where:

Note: k*L depends on the reaction rate, not the hydrodynamics which affects k L, and CA 0 for these fast reactions since the gas reacts as soon as it is transferred. The ratio of the diffusion to reaction times in the neighborhood of the gas liquid interface is conveniently expressed as (tD/tR) which is also known as the Hatta Number (Ha). Its value determines the extent to which diffusion controls the liquid phase reaction. Several regimes are identified by the range of (tD/tR) values and each regime has different requirements for reactor design (see Table 1). Resistance to diffusion in the gas phase has to be accounted for in many cases. The rate of transfer through the gas phase alone is given by:



Where: kG = gas film mass transfer coefficient PA = partial pressure of A in the gas phase p*A = partial pressure of A at the interface.

Generally more than one resistance is important, so they have to be combined into an overall coefficient KL, e.g.:

Where: He = gas-liquid equilibrium constant

= Henry's Law coefficient, i.e.:

The regimes can be seen, for example for n = 1 and m = 1, on a graph of the enhancement factor k*L/kL against the Hatta Number, given in Figure 2.



FIGURE 2 ENHANCEMENT FACTOR vs HATTA NUMBER (for Case m = 1, n = 1)

Note: The group is the ratio of the liquid bulk to film volumes. Table 1 lists the Regimes (marked I to V in Figure 2), together with the conditions, important variables, diagrams of the concentration profiles and the local rate equations.



TABLE 1 REGIMES OF GAS-LIQUID MASS TRANSFER WITH ISOTHERMAL CHEMICAL REACTION



6.3 Modification for Exothermic and Complex Reactions Exothermic reactions occurring in the film or interface can dramatically influence the enhancement factor. The rise of temperature near the interface affects the gas solubility (decreasing the reaction rate), but also in general increases the reaction rate (Arrhenius effect). The juxtaposition of these effects can lead to multiplicity of steady states. This is illustrated in Figure 3 (from Ref. [23] and Ref. [1]) for SO3 reacting with dodecylbenzene.



The following list covers complexities of reactions which have been covered in the literature: (a) Simultaneous Absorption of two Reacting Gases

An example is the absorption of H2S and CO2 in ethanolamines.

See Cornelissen (Ref. [13]), summarized in Charpentier (Ref. [11]), demonstrating resolution of the film model equations and the possibility of apparently negative kG. Also see Ref. [12].

(b) Reactive Particles within the Liquid Film (c) Autocatalytic Reactions in the Film (see Ref. [37]) (d) Reaction Forming Volatile Products

These can occur if the liquid diffusivity of the product is less than that of the reactant. See Ref. [11] and Ref. [36].

(e) Reversible Reactions

See Ref. [13], Ref. [11], Ref. [16] and Ref. [10].

(f) Free Radical Multiplying Chain Reactions See Ref. [18].

(g) Consecutive Reactions in the Film

See Ref. [8].



FIGURE 3 ENHANCEMENT FACTOR vs HATTA NUMBER : EFFECT OF THERMAL & OTHER FACTORS



6.4 Preliminary Selection of Reactor Type Table 1 shows how the various factors dominate reactor behavior according to the regime. A preliminary choice of the reactor type can therefore be made to give designs (e.g. film or bulk liquid, high or low kL, etc.) appropriate to the regime expected, having regard for practically achievable kG, kL, a and eL values. Table 2 presents suggestions for cases without large exotherms or other complications. Before a final choice can be made other considerations intervene, e.g.: (a) The overall (macro) flow pattern; (b) Heat transfer requirements; (c) Materials of construction; (d) Foaming properties. These are dealt with in Clause 9. Table 3 provides a summary of comparative mass transfer performance estimates.



TABLE 2 REGIMES OF GAS-LIQUID MASS TRANSFER IGNORING LARGE EXOTHERMS OR OTHER COMPLICATIONS



7 EXPERIMENTAL DETERMINATION OF REGIME There are two basic methods to determine the mass transfer/reaction regime and thence select and design or scale-up the reactor. One method consists of measuring the true liquid phase reaction kinetics (without mass transfer interference) directly and then using this information with typical values of the mass transfer coefficient (e.g. from Table 3) for the candidate reactors selected from Table 1. When this is not possible, the other method is to run the reaction in a laboratory-scale gas-liquid reactor and follow its responses to certain changes of conditions. 7.1 Direct Measurement of Reaction Kinetics The direct measurement of the reaction kinetics is difficult for fast reactions since it is necessary to remove all mass transfer influences, i.e. to work in Regime I. With slow reactions a bubble column or stirred vessel will operate in Regime I. With somewhat faster reactions a more intense contactor (i.e. high kL) might be used. Examples are the turbulent static mixer (see 11.11) and the spinning film reactor (see 11.7). In either case, experiments should preferably be carried out under steady state conditions and as nearly at the intended plant operating temperature, pressure and concentrations as possible with the rate measured by mass balance between the inlet and outlet streams. The residence time of the liquid phase should be measured either by tracer tests or by measuring the liquid throughput and the liquid holdup volume in the reactor. Prior to carrying out the kinetic measurements the absence of any effect due to changing gas rate or stirrer speed on the measured rate of reaction per unit liquid volume should be confirmed. If laboratory-scale intense contactors still show a mass transfer influence, the reactions are fast and mass transfer will influence the performance of the plant-scale gas-liquid reactor (Regimes II - V). It may be possible to predissolve the gas in the reaction solvent (probably at different conditions from the plant operating conditions), in which case mass transfer effects are eliminated (assuming no gas desorbs) and one of the following liquid phase techniques may be used (Figure 4 illustrates the first two):



(a) Stopped Flow Technique

This technique is currently limited to < 70°C, atmospheric pressure and low viscosities in glass, followed by spectrometric analysis. Ninety-nine percent completion of the reaction should occur in not more than 5 milliseconds (see Ref. [7]).

(b) Continuous Flow T-jet mixed Reactor

This technique is suitable for higher pressures and temperatures, can be made from various materials of construction, but is limited to low viscosity liquids and reaction times tR >> 10 milliseconds. Analysis may be by temperature rise (should only be a few degrees), by spectrophotometer (using windows) or from samples (rapid quenching is required). See Ref. [33].

(c) Temperature Jump

The reactants are mixed at a low temperature (such that the reaction rate is negligible) and the temperature is instantly raised to the desired reaction condition and progress of the batch is monitored (see Ref. [3]).

(d) Pressure Jump

This technique is similar to the temperature jump but is intended for gas-liquid reactions (i.e. without predissolving).



FIGURE 4 REACTORS FOR LIQUID-PHASE KINETICS MEASUREMENT

7.2 Laboratory Gas-Liquid Reactor Experiments 7.2.1 Special Laboratory Devices

A number of specialized devices have been developed and used for determining which regime applies. The idea is that they provide a known interfacial area and mass transfer coefficients kL and kG which are the same as or can be related to those at a point in a plant scale reactor. These devices have been reviewed by Charpentier (Ref. [11]). Table 3 summarizes the devices and their characteristics.



The spinning cone reactor could be added to the list with a time of contact between 0.01 and 0.15 seconds, kL of about 0.5 cm/s and an interfacial area of 50 to 300 cm2. Comparison with Table 4 will indicate which devices are most likely to represent a given plant contactor. The experiments should cover the range of concentrations and partial pressures expected in the plant reactor.

TABLE 3 COMPARATIVE MASS TRANSFER PERFORMANCE OF

CONTACTING DEVICES



7.2.2 Laboratory Stirred Reactors Laboratory stirred reactors are more commonly available than the specialized devices and tend to serve as both "kinetic machines" and "semi-tech reactors". If designed according to the standard configuration given in GBHE-PEG-MIX-705, the values of kLa and gas hold-up determined in the laboratory can be scaled up. Laboratory stirred reactors can therefore give useful design information for all regimes so long as the same regime operates throughout the vessel (beware that many laboratory autoclaves have a non-standard configuration and use surface entrainment of the gas, which affects the mechanism, and hence model, since the gas feed rate depends on the agitator speed). The partial backmixing of the gas phase may however lead to uncertainties in relating gas concentration profiles. A thorough study of the film reactions is best carried out in one of the film reactors described in 7.2.1 and Table 3. The stirred reactor should be set up to operate at conditions as near as possible to those intended for the full-scale plant when producing the desired products and yields. No materials should be omitted if they are to be present in the final process. The rig should be so designed that agitator speed, temperature, pressure and gas concentration can be varied. Continuous steady-state experiments are recommended. Figure 5 proposes a sequence of experiments to establish which regime is operative and hence guide in the selection of an appropriate reactor type. Each change of condition should be approached from both directions to confirm that regime "boundaries" have not been crossed in the process. Note: Processes during which solvent or product boils off may not respond in the ways inferred by Figure 5 because changes in reaction rate alter the gas hold-up (and hence the volume of liquid in the reactor if it is level-controlled) and the gas phase reactant concentration. Figure 6 is an extension of Figure 5 to incorporate cases in which a solid catalyst in suspension participates in the reaction.



8 EQUILIBRIUM AND DIFFUSIVITY DATA SOURCES The use of laboratory stirred reactor experiments and the scale-up methods described in Clause! 9 eliminate the need for data on gas-liquid equilibria (Henry's Law coefficient, He) and diffusivity (DLA) in the design. These will, however, be required if the design uses the basic reaction rate constants and the previously determined mass transfer coefficients. Some data will be found in Perry's "Chemical Engineer's Handbook" and in International Critical Tables. Data on gas solubility can also be obtained from standard collections such as Ref. [38], Ref. [35], Ref. [26] and Ref. [4]. For simple molecules the solubility can be estimated from the surface tension, molecular dimensions and interaction energy (see Ref. [39]). Chemical Abstracts could also be consulted. FIGURE 5 EXPERIMENTS TO DETERMINE THE OPERATING REGIME



FIGURE 6 EXPERIMENTS TO DETERMINE THE OPERATING REGIME WHERE A SOLID CATALYST IS INVOLVED

9 OVERALL EFFECTS

So far only the "local" effects on the reactions have been described. It is important to consider how local conditions change throughout the reactor and what effects the reaction has on the design of the reactor as a whole.

9.1 Liquid Flow Patterns

A reaction in Regime I 'sees' the flow patterns in the liquid phase, and may be influenced by the resulting liquid phase concentration changes through the reactor, as outlined below. Each reactor type gives a different flow pattern: see Clause 10.

Three types of ”axial mixing” (mixing in the general flow direction or ”residence time distribution”) are distinguishable for ”ideal” reactors:



(a) Batch: All reactant is in the reactor for the same period. Mean concentration = log mean of initial and final.

(b) Plug flow: As in 9.1(a) but continuous. Tubular reactors in

turbulent flow, or several stirred vessels in series, approximate to this.

(c) Backmixed: Reactant has a wide distribution of residence times in

the reactor; mean concentration = outlet concentration. Continuous stirred vessels, or plug flow reactors with large recycle loop, aim to approximate to this.

These types should not be confused with ”radial” or ”cross” mixing, which is uniformly across the general flow direction and is required for all reactors. For reactions of the order n 1, the required reactor volume will evidently be greater for a Backmixed reactor than for plug flow of liquid (e.g. by a factor of 3.91 for 90% conversion, 21.5 for 99% conversion, etc. for first order reaction). This is also true for gas-liquid mass transfer controlled systems, where it is the gas phase axial mixing which affects the driving force and hence reactor size for a given throughput. For certain more complicated reactions, the yield of the reaction is also influenced by the mode of mixing. Full discussion of this is outside the scope of this Guide (see Ref. [14] for an introduction), but, briefly, for consecutive reactions: A B C high yield of B (compared with C) is favored by plug flow, but if C is the desired product, a backmixed reactor should be chosen. For parallel reactions: A B (product) (of order p in A) A C (of order q in A) if p > q, high yield of B requires high CA, hence plug flow is required, conversely if p < q, high CA has to be avoided, so a backmixed reactor is preferred. If p = q, the yield is unaffected by the mixing mode.



9.2 Scale of Mixing To see the above effects, obviously the timescale of mixing has to compare with the reaction time. The backmixing found in stirred vessels etc. may be described by a seconds minutes timescale (often called ”macromixing”). If, however, the reaction is mixing sensitive and is substantially complete within milliseconds, any yield effects will arise from mixing on the microsecond millisecond timescale (”micromixing”). In such cases local effects round the reactant feedpoint (or round the bubbles if a gas reactant is involved) are controlling, and overall residence time distribution is irrelevant to the reaction selectivity. High intensity mixers (e.g. turbulent static mixers) should be considered here. Bourne (Ref. [6]) discussed modelling of micromixing. 9.3 Gas Flow Pattern: Mean Driving Force for Mass Transfer In regimes other than I, the overall axial mixing (residence time distribution) of the gas is important, as it affects the mean concentration difference which drives the gas-liquid mass transfer. Gas flow patterns depend upon the type of reactor (see Clause 4) and sometimes its geometry and gas velocity. In using equation (2) over the whole reactor, the appropriate mean value of the ‘driving force' has to be used. Since CA y 0, for batch gas (rare) or ideal plug flow of gas, this mean value



and for fully backmixed gas, the mean value Hence backmixing gives lower mean driving force (implying larger reactor volume), but lower peak gas concentration (which may influence the reaction selectivity) than plug flow or batch gas. In real reactors the ideal flow patterns are not achieved, but may be approached. More detailed flow modeling is required if deviations from the ideals are substantial, in other words where selectivity is governed by reactions of timescale similar or faster than the mixing timescale. 9.4 Gas-Liquid Reactor Modeling More complicated approaches are required when the reaction yield is sensitive to the mode of mixing, and is carried out in a reactor which is neither in ideal plug flow, nor perfectly mixed, but somewhere between. The mixing is then described by a model of the flow pattern, whose equations have to be solved together with those describing the reaction kinetics. Since backmixing implies some degree of recirculation, the solution of the equations has to be iterative and generally requires a substantial computer program. There are three types of mixing model for continuous reactors, based on: (a) Arbitrary zones of mixedness and segregation (or dispersion and

recoalescence) of elements of fluid; (b) Measured flow patterns; (c) Calculated flow patterns. Much effort has been put into developing type (a) models, (see Ref. [31]) but they remain unsuitable for scale-up use because they require fitted parameters which are not available. In other words they can describe mixing but not predict it. They have not been applied to gas-liquid systems.



Type (b) models are at present the most useful, though still requiring more basic data. Mann (Ref. [23]) describes a successful example, with models of the liquid and gas circulations in a stirred vessels as loops of x connected mixed zones (see Figure 7) with feed and offtake where appropriate. The zones can have different kLa and gas hold-up if these are determined experimentally. The mean circulation flow and relevant number of zones, x, have to be supplied from experiment, so the model is again not in itself predictive, also the parameters have physical meaning, so are more amenable to scale-up. Middleton (Ref. [27]) describes a measurement technique and some scale-up correlations (applicable to zero or low gas holdups, with the standard vessel configuration) for mean circulation time

and circulation time distribution, which can be converted to the number of zones, x, in a loop.

A more detailed modelling technique using the same information is proposed in Ref. [25].Patterson (Ref. [29]) describes a reactor model which uses actual local fluid velocity and turbulence measurements throughout a vessel to describe the flow pattern (simplified to a rectangular grid of 'cells'). This is getting even nearer the true hydrodynamics, but does not yet have sufficient velocity data behind it for reliable application to large vessels or gas-liquidsystems. The ultimate model is type (c) which involves a complete solution of the 'Navier Stokes' equations (with a turbulence model), together with the reaction equations, using 'Computational Fluid Dynamics' (CFD). Middleton (Ref. [28]) gives an example of this for competing reactions in turbulent stirred reactors. A wide variety of configurations and conditions can be modelled in detail. GBHE-PEG-RXT-800 should be consulted for guidance to experts.



FIGURE 7 THE MIXED ZONES IN LOOPS' MODEL FOR STIRRED REACTORS

9.5 Heat Transfer Obviously the reactor has to be capable of transferring (supplying or removing) heat at the appropriate rate, which is a question of supplying sufficient boil-off or condensation rate, or if sensible heat is being transferred, sufficient temperature difference, heat transfer area, and liquid velocities over the transfer surfaces. Design correlations for liquids are widely available, but less so for gas-liquid systems. Correlations are available (see GBHE-PEG-MIX-705) for stirred vessels and bubble columns, which are frequently used where heat transfer is a problem. For low gas volume fractions (at least in stirred vessels), say 15%, Edney and Edwards (Ref. [17]) show that the gas does not significantly affect the heat transfer to the liquid. With typical boiling systems the extra bubbles



affect the hydrodynamics, increasing backmixing in bubble columns, and adding to gas 'flooding' of agitators (for safe design use the maximum, i.e. exit, gas rate in hydrodynamic calculations). Gas absorption 'driving force' is also reduced by the dilution of the gas phase by vapour. High boil-off vapour velocities (say > 0.05 m/sec) should be avoided to prevent carry-over of liquid mist into the vapour off-take. 9.6 Materials of Construction Selection of materials of construction for plant reactors is often not easy and should be made at an early stage of experimentation to ensure that research work is done under conditions, and in devices, that are practical at full scale considering the materials to be used. Conditions in reactors are often corrosive to common materials; in gas-liquid reactors erosion/cavitation phenomena add to these problems, especially for spargers, agitator blades, baffles, and other high-velocity zones. Specialist engineers should be consulted if problems are anticipated. High intensity devices (tubular jet mixers, turbulent static mixers), where applicable, provide small materials weight per unit performance and so may allow more exotic materials to be used than with vessels (although erosion problems may be worse). 9.7 Foaming Foaming should be anticipated in all gas-liquid reaction cases unless known to be absent, since it is not possible to predict whether a foaming problem will occur. Surface active solutes, fine non-wetter solid particles, rapid gas-liquid mass transfer and reactions near the gas-liquid interface can all lead to foaming problems. Methods of dealing with foaming are reviewed in GBHE-PEG-MIX-705 and in the GBH Enterprises Mixing and Agitation Manual. 10 FINAL CHOICE OF REACTOR TYPE Tables 4 and 5 summarize the typical properties of gas-liquid reactors from which the selection may be made. The values given in Table 4 are comparative and should not be used for design calculations. The design values will depend on the physical properties of the fluids in the reactor.



TABLE 4 COMPARATIVE MASS TRANSFER DATA (low viscosity systems, non-boiling)

Note: These are orders of magnitude for comparison, not design figures. They are based on air and water physical properties at 20 °C. Approximate adjustment for other systems may be made by:

(Calderbank, Ref. [9]), though others find the index of DAL to be 1. For further background see Ref. [32], Ref. [22] and Ref. [21].



TABLE 5 CHOICE OF GAS-LIQUID REACTOR TYPE



11 SCALE-UP AND SPECIFICATION OF GAS-LIQUID REACTORS This Clause gives some sources of reference to correlations which can be used in the design of gas-liquid reactors or in the scale-up from laboratory rigs or other plants. Depending on the regime, the design will require quantitative information on: (a) Practical range of reactor sizes. (b) Range of practical flowrates. (c) Gas hold-up (voidage fraction). (d) Interfacial area per unit volume. (e) Mass transfer coefficient (often together with (d) above):

(1) KL for liquid-side diffusion control. (2) KG for gas-side diffusion control.

(f) Liquid backmixing. (g) Gas backmixing. (h) Heat transfer. 11.1 Bubble Columns Bubble columns are simple to build and maintain, have no rotating seal problems and can be made in large sizes. However, they have a low mass transfer intensity (especially for coalescing systems), an uncontrolled degree of liquid backmixing and mediocre heat transfer performance. Low concentrations of easily-suspended solids can be handled. For scale-up correlations consult GBHE-PEG-MIX-705. 11.2 Packed Columns These liquid film devices have the advantage of counter-current flow (where applicable, i.e. for incomplete absorption from a multicomponent gas stream), low pressure drop and a wide range of available materials of construction. They can be designed for high or low gas/liquid ratios. The flowrates are limited by hydraulics. For design for mass transfer consult GBHE-PEG-MAS-612, manufacturers' literature (e.g. Norton, ETA, Sulzer) and/or Ref. [32].



11.3 Trickle Beds These are co-current downflow, low intensity variants of the packed column and are often used where the packing is or supports a solid catalyst and where co-current flow has no disadvantage (e.g. irreversible simple reactions with complete consumptions of adsorbed components or when the gas phase is a single component). Pressure drop is less and hydraulic limits are wider than those for countercurrent columns. Heat transfer is difficult and poor. Design information can be found in: (a) Froment (1979), page 70 (Ref. [19]). (b) Herskowitz (1983) (Ref. [20]). (c) Satterfield (1975) (Ref. [34]). 11.4 Plate or Tray Columns Plate or tray columns are similar to packed columns, but have a larger liquid hold-up. The hydraulic operating range is wider and heat transfer is easier. However blockages due to the presence of solids and problems due to foaming may be more severe. The mass transfer and hydraulic design is usually based on the methods developed by Fractionation Research Inc. (e.g. their program FRISIEVE for sieve-plate columns). GBHE-PEG-MAS-611and Ref. [32] should also be consulted. 11.5 Spray Columns Spray columns are rarely used because coalescence of droplets often results in a mediocre interfacial area, and the entrainment of fine droplets into the gas outflow can be a problem. They are, however, simple and easily cleanable (therefore suitable for "dirty" gases) and have a low pressure drop on the gas side. Design information is sparse; see Froment (1979), page 725 (Ref. [19]).



11.6 Wiped Film Wiped film devices are commonly sold as evaporators for heat sensitive or high viscosity liquids. Mass transfer intensity is high, but such devices require a heavy maintenance effort. For design information consult the manufacturers (e.g. Luwa). 11.7 Spinning Film Reactors These devices are still at the research stage as reactors, although centrifugal thin-film evaporators are available commercially (e.g. from Alfa-Laval and Luwa). With a 45° cone, gas-liquid mass transfer is enhanced by surface waves on the film giving rise to very high values of KL for low viscosity liquids. For more information, consult the experts listed in GBHE-PEG-RXT-800. 11.8 Stirred Vessels Stirred vessels are widely used (in the higher KL regimes) as they provide flexibility, good intensity of mass and heat transfer and some control of bubble size and backmixing. The additional cost of the agitator system and possible shaft sealing difficulties should, however, be considered compared with the simplicity of a bubble column. Stirred vessels should be used where there are solids suspension problems. Because of shaft mechanical limitations they are not usually very tall. Design correlations are given in GBHE-PEG-MIX-705. 11.9 Plunging Jet The plunging jet is a low-cost low-intensity contactor which has recently come into use for aerating effluent ponds and reservoirs. The energy efficiency is good. There is little design information available, but see Ref. [5].



11.10 Surface Aerator The surface aerator is the older, mechanical equivalent of the plunging jet. It is less energy efficient and requires more maintenance. The degree of mixing is difficult to quantify. Design information can be obtained from manufacturers (e.g. SEM, Lightnin). 11.11 Static Mixers When used with low viscosity liquids in the turbulent regime, static mixers give very high intensities of gas-liquid mass transfer and approach plug flow conditions. They offer good heat transfer as they can be used within shell-and-tube heat exchangers. They can be operated at low gas/liquid ratios in bubble flow for Regime II or at high gas/liquid ratios and gas velocities in film flow (liquid film on all internal surfaces) for Regimes IV and V or under gas-film controlled situations. Static mixers are new in the gas-liquid reaction area but there are some plant applications. "Honeycomb" type static mixers can also be useful for reactions between a gas and a high viscosity liquid. Design data may be obtained from GBHE-PEG-MIX-705, from Section E4 of the GBHE Mixing and Agitation Manual or from Ref. [30]. 11.12 Ejectors, Venturis and Orifice Plates Ejectors, venturis and orifice plates can provide high intensity mass transfer if correctly designed. Ejectors are used in the Buss loop reactor which has several commercial applications (e.g. H-Acid plant, FCMO). Ejectors and venturis are covered in GBHE-PEG-MIX-705. Orifices are a crude alternative and are not generally recommended because of the separated flow zones behind the orifice plates where coalescence and extended residence times may occur.



11.13 3-Phase Fluidised Bed The upward flow of gas and liquid supports a bed of particles which can enhance break-up or coalescence of bubbles depending on the properties of the particles. They are not widely used or understood. Limited design information is given in GBHE-PEG-MIX-706 and Section F2 of the GBH Enterprises Mixing and Agitation Manual.



12 BIBLIOGRAPHY [1] Allan J C & Mann R (1979), Chem. Engineering Science, 34, 413. [2] Astarita (1967), "Mass Transfer with Chemical Reaction", Elsevier. [3] Bamford C H (1969) & Tipper C F H (eds), "Comprehensive Chemical Kinetics", Vol. 1, p.134 (Elsevier). [4] Battino (1966) & Clever, Chem. Rev., 66, 395. [5] Bin A K (1982) & Smith J M, Chem. Engineering Commun. 15, 367. [6] Bourne J R (1982), Chem. Engineering Commun, 16, 79. [8] Brian P L T (1965) & Beaverstock M C, Chem. Engineering Science, 20, 47. [9] Calderbank P H (1961) & Moo-Young M B, Chem. Engineering Science, 16, 39. [10] Carmichael G R (1980) & Chang S C, Chem. Engineering Science, 35, 2459 and 2463. [11] Charpentier J C (1982), Trans.IChemE., 60, 131. [12] Chaudhari R V (1975) et al, Chem. Engineering Science, 30, 945. [13] Cornelissen A E (1980), Trans.IChemE., 56, 242. [14] Coulson J M (1971) & Richardson J F, "Chemical Engineering", Vol. 3, Second edition, Pergamon Press. [15] Danckwerts P V (1970), "Gas-Liquid Reactions", McGraw-Hill. [16] deCoursey W J (1982), Chem. Engineering Science, 37, 1483. [17] Edney H G (1976) & Edwards M F, Trans. IChemE, 54, 160. [18] Flores-Fernandez G (1978) & Mann R, Chem. Engineering Science, 33, 1545. [19] Froment G F (1979) & Bischoff K B, "Chemical Reactor Analysis and Design", Wiley.



[20] Herskowitz M (1983) & Smith J M, AIChE J., 29, 1. [21] Horak J (1978) & Pasek J, "Design of Industrial Chemical Reactors from Laboratory Data", Heyden. [22] Laurent A (1983) & Charpentier J C, Int. Chem. Engineering, 23, 265. [23] Mann R (1977) & Moyes H, AIChE J. 23, 17. [24] Mann R (1977a), Middleton J C & Parker I B, Proc. 2nd Eur. Conf. Mixing, BHRA, Cambridge, paper F3. [25] Mann R (1981), Manros P P & Middleton J C, Trans.IChemE, 59, 271. [26] Markham A (1941) & Kobe K, Chem. Rev., 28, 519. [27] Middleton J C (1979), Proc. 3rd Eur. Conf. Mixing, BHRA, York, paper A2. [28] Middleton J C (1986), Pierce F & Lynch P M, ChERD, 64, 18. [29] Patterson G K (1975), in Brodkey (ed) "Turbulence in Mixing Operations", Academic Press. [31] Rao D P (1973) & Edwards L L, Chem. Engineering Science, 28, 1179. [32] Rase H F (1977), "Chemical Reactor Design for Process Plants", Wiley. [34] Satterfield C N (1975), AIChE J., 21, 209. [35] Seidell A (1965), rev. Linke W, "Solubilities", American [36] Shah Y T (1974), Pangarkar V G & Sharma M M, Chem. Engineering Science, 29, 1601. [37] Sim M T (1975) & Mann R, Chem. Engineering Science, 30, 1215. [38] Stephen H (1963) & Stephen T (eds), "Solubilities of Inorganic and Organic Compounds", Pergamon. [39] Uhlig H H (1937), J phys. Chem., 41, 1215.



DOCUMENTS REFERRED TO IN THIS PROCESS ENGINEERING GUIDE This Process Engineering Guide makes reference to the following documents: GBHE ENGINEERING GUIDES Glossary of Engineering Terms (referred to in Clause 3) GBHE-PEG-MAS-611 Design and Rating of Trayed Distillation Columns (referred to in 11.4). GBHE-PEG-MAS-612 Design and Rating of Packed Distillation Columns (referred to in 11.2). GBHE-PEG-MIX-705 Mixing of Gas Liquid Systems (referred to in 7.2.2, 9.5, 9.7, 11.1, 11.8, 11.11 and 11.12). GBHE-PEG-MIX-706 Gas-Solid-Liquid Mixing Systems (referred to in 11.13). GBHE-PEG-RXT-800 Experts on Reactor Technology (referred to in 9.4 and 11.7). GBHE-PEG-RXT-809 Homogeneous Reactors (referred to in 6.1). OTHER GBHE DOCUMENT GBHE Mixing and Agitation Manual (referred to in 9.7, 11.11 and 11.13).

